Configurations and methods for deep feed gas hydrocarbon dewpointing

ABSTRACT

An natural gas processing plant allows for recovery of at least 98% of butane and heavier hydrocarbons, and about 60 to 80% of propane hydrocarbons from a rich feed gas stream with a single fractionator that operates at two different pressures, that receives a chilled gas from a turboexpander in the upper fractionator and a C5+ liquid in the lower section, while producing a C2− vapor stream in the lower section that is used as reflux to the upper section. Most typically, contemplated configurations and methods operate without the use of external refrigeration.

FIELD OF THE INVENTION

The field of the invention is removal and recovery of natural gas liquids (NGL) from feed gases to meet pipeline hydrocarbon dew point and heating value specifications.

BACKGROUND OF THE INVENTION

Numerous systems and methods are known in the art to recover C2, C3, and heavier components from natural gas, but all or almost all of them are configured for high recovery (i.e., over 90%) of NGL and require use of a turboexpander and deep refrigeration, which are costly and can be only be economically justified if there are significant downstream markets. With increasing demands on the condensate portion (i.e., C5 and heavier) and less demand for LPG (i.e., C2, C3, and C4) components, the high capital investment and operating costs required for high recovery can typically not be justified. On the other hand, pipeline operators are required to produce a sales gas to meet the pipeline specification on the hydrocarbon dew point and heating value for safety in transmission. In most cases, recovery of over 98% of the C4 and heavier hydrocarbons from the feed gas is required, while C3 recovery can be as low as 60%. In view of the changed demand, the complexity of currently known NGL processing plants that allow over 90% C3 recovery is excessive and can often not be justified from an economical perspective.

For example, numerous NGL processing plants with high NGL recovery from a feed gas include cryogenic fractionation and turbo-expansion processes as described in U.S. Pat. No. 4,157,904 to Campbell et al., U.S. Pat. No. 4,251,249 to Gulsby, U.S. Pat. No. 4,617,039 to Buck, U.S. Pat. No. 4,690,702 to Paradowski et al., U.S. Pat. No. 5,275,005 to Campbell et al., U.S. Pat. No. 5,799,507 to Wilkinson et al., and U.S. Pat. No. 5,890,378 to Rambo et al., and U.S. Pat. App. No. 2002/0166336 to Wilkinson et al., and WO 2011/126710 to Johnke et al. These and all other extrinsic materials discussed herein are incorporated by reference in their entirety. Where a definition or use of a term in an incorporated reference is inconsistent or contrary to the definition of that term provided herein, the definition of that term provided herein applies and the definition of that term in the reference does not apply.

However, while all of these processes can achieve very high NGL recovery, several difficulties still remain. Among other things, the NGL recovery processes use high expansion ratio turboexpanders to produce low levels of refrigeration, which requires recompression of the residue gas. Moreover, when processing a rich gas stream with relatively high levels of C5+ hydrocarbons, additional external refrigeration is often required. Typically, such process configurations are complex and are difficult to operate. For example, Campbell et al. describe in U.S. Pat. No. 6,182,469 a plant in which feed gas is cooled in a heat exchanger using cold residue gas and side reboilers as depicted in Prior Art FIG. 1. The condensed feed gas liquids are then separated in a separator and fed to the demethanizer. Alternatively, as described by Sorensen in U.S. Pat. No. 5,953,935, an absorber may be added upstream of a demethanizer as depicted in Prior Art FIG. 2. In such configurations, the liquids from the feed separator and the absorber bottoms are fed to the demethanizer To further increase NGL recovery in such configurations, the absorber overhead is cooled and refluxed by chilling with the demethanizer overhead vapor.

In still further known configurations, as described in U.S. Pat. No. 6,244,070 to Lee et al. and U.S. Pat. No. 5,890,377 to Foglietta, the reboiler duties are integrated in feed chilling, and in these configurations, liquids from the intermediate separators are fed to various positions in the downstream demethanizer for NGL recovery. These processes also include various means of providing cooling to the NGL processes. Exemplary known configurations following such schemes are depicted in Prior Art FIGS. 3 and 4.

While such complex configurations are suited to achieve high C2 and C3 recovery to over 95%, they tend to be not cost-effective for C4+ and moderate C3 recovery (e.g., 60-75%) due to their relatively high expansion ratio and energy demand for added refrigeration. Therefore, although various configurations and methods are known to recover NGL from a feed gas, all or almost all of them suffer from one or more disadvantages when moderate C3 recovery is required. Therefore, there is still a need to provide methods and configurations for improved NGL recovery.

SUMMARY OF THE INVENTION

The inventive subject matter is directed to configurations and methods of recovery of C4 and heavier hydrocarbons, and moderate recovery (up to 75%) of C3 from a gas stream to meet hydrocarbon dew point and heating value specification of a pipeline gas produced from the gas stream.

In one preferred aspect of the inventive subject matter a method of hydrocarbon dew point adjustment of a natural gas that includes C3 and C4 and heavier components has a step of cooling the feed gas in a feed gas exchanger using a liquid phase of the cooled feed gas and an overhead product of an upper section of a fractionator. In another step, the cooled feed gas is separated in a phase separator into the liquid phase and a vapor phase, and the liquid phase is fed into a lower section of a fractionator while the vapor phase is fed into the upper section of the fractionator. In especially preferred aspects, the upper and lower sections of the fractionator are coupled to each other such that an expansion device (typically a JT valve) reduces pressure of and provides a vapor product of the lower section to the upper section, and such that a pump increases pressure of and provides a liquid product of the upper section to the lower section. In yet another step, the vapor product of the lower section is cooled in an overhead exchanger using refrigeration content in the overhead product of the upper section of the fractionator. In such methods, it is generally preferred that the fractionator is operated such that C3 recovery from the feed gas is between 60% and 80%, and recovery of the C4 and heavier components from the feed gas is at least 95%.

It is further generally preferred that the fractionator is operated at a pressure of between 450 to 550 psig, and that the upper section is operated at a pressure that is at least 10 psig, and more typically at least 30 psig lower than the pressure of the lower section. With respect to temperatures, it is typically preferred that the upper section is operated at a temperature of −65° F. to −55° F. and that the lower section is operated at a temperature of 25° F. to 300° F.

Additionally, contemplated methods will also include a step of expanding the vapor phase in a turbo expander and reducing pressure of the liquid phase in a second expansion device before feeding the vapor phase and the liquid phase into the upper and lower sections of the fractionator, respectively. While not limiting to the inventive subject matter, it is typically preferred that the feed gas cooling and/or cooling of the vapor product of the lower section is performed without use of external refrigeration.

In another preferred aspect of the inventive subject matter, a processing plant for hydrocarbon dew point control of a natural gas feed gas delivered from a feed gas source will include a feed gas exchanger that is fluidly coupled to the feed gas source and configured to cool the feed gas using a liquid phase of the cooled feed gas and an overhead product of an upper section of a fractionator. Contemplated plants will also include a phase separator that is fluidly coupled to the feed gas exchanger and that is configured to separate the cooled feed gas into the liquid phase and a vapor phase. Most typically, the fractionator comprises a lower section that is configured to receive the liquid phase and an upper section configured to receive the vapor phase in the upper section. Especially preferred fractionators have upper and lower sections coupled to each other such that an expansion device reduces pressure of and provides a vapor product of the lower section to the upper section, and such that a pump increases pressure of and provides a liquid product of the upper section to the lower section. An overhead exchanger is included and configured to cool the vapor product using the overhead product of the upper section of the fractionator.

In particularly preferred plants, a second expansion device that is included and configured to receive and reduce pressure of the liquid phase, and a turbo expander included and configured to receive and reduce pressure of the vapor phase. It is especially contemplated that the pump increases the pressure of the liquid product in an amount of at least 10 psig, and more typically at least 30 psig, and that the fractionator is configured to operate at a pressure of between 450 to 550 psig. Thus, the upper section and the lower section of the fractionator are configured to operate at a pressure differential of at least 10 psig. It is further especially preferred that the upper section of the fractionator is configured to operate at a temperature of −65° F. to −55° F. and wherein the lower section of the fractionator is configured to operate at a temperature of 25° F. to 300° F., and/or that the feed gas exchanger, the fractionator, and the expansion device are configured to allow for between 60% and 80% C3 recovery without use of external refrigeration.

Various objects, features, aspects and advantages of the present invention will become apparent from the following detailed description of preferred embodiments of the invention, along with the accompanying drawings.

BRIEF DESCRIPTION OF THE DRAWINGS

Prior Art FIG. 1 is a schematic of one known configuration for NGL recovery in which feed gas is cooled in a heat exchanger using cold residue gas and side reboilers.

Prior Art FIG. 2 is a schematic of another known configuration for NGL recovery in which an absorber/fractionator column is positioned upstream of a demethanizer

Prior Art FIG. 3 is a schematic of yet another known configuration for NGL recovery in which reboiler and feed gas compression are integrated in feed chilling.

Prior Art FIG. 4 is a schematic of a further known configuration for NGL recovery in which reboiler and compressed residue gas recycle are integrated in feed chilling.

FIG. 5 is a schematic of an exemplary configuration for NGL recovery according to the inventive subject matter.

FIG. 6 is a table listing calculated compositions of gas streams in the exemplary NGL recovery plant of FIG. 5.

DETAILED DESCRIPTION

The inventor has discovered various configurations and methods of NGL recovery in which capital and operating cost are significantly reduced, especially where a rich feed gas is processed and where C4+ recovery with moderate C3 recovery is required. Among other advantages, contemplated configurations and methods significantly reduce complexity and cost by reducing the number of equipment services, by elimination of external refrigeration, and lowering residue gas compression requirements.

In particularly preferred configurations and methods, the feed gas (typically a natural gas comprising C3, and C4 and heavier components) is cooled at relatively high pressure to thereby effect partial condensation. The vapor and liquid phases are then separated, with the liquid phase being expanded to a lower pressure to so provide cooling to the feed gas. After reduction in pressure, the liquid phase is fed to the lower section of a fractionation column, while the vapor phase is expanded via a turboexpander and fed into the upper section of the fractionator. As the fractionator is operated at relatively high pressures (typically 450 to 550 psig), recompression requirements are significantly reduced.

Moreover, it is preferred that the fractionator is a single fractionation column that has at least two different pressure sections and a booster pump that is fluidly coupled between the sections to provide for the pressure differential and to pump liquid from the upper section (which is lower in pressure relative to the lower section) to the lower section, which produces a methane- and ethane-rich vapor stream. In still further preferred aspects, the methane- and ethane-rich vapor stream is cooled and partially condensed by the fractionator overhead and then fed to the top section as the reflux. It should thus be appreciated that such processes will significantly lower capital and operating costs, particularly where 98% of C4 and heavier hydrocarbons and moderate C3 recovery, typically 60% to 80%, are required. As used herein, and unless the context dictates otherwise, the term “coupled to” is intended to include both direct coupling (in which two elements that are coupled to each other contact each other) and indirect coupling (in which at least one additional element is located between the two elements). Therefore, the terms “coupled to” and “coupled with” are used synonymously.

Contemplated configurations and methods are particularly advantageous in processing a relatively rich gas stream (e.g., at least 3% C3, and at least 2.5% C4+), in which the feed gas is cooled to allow removal of at least some of C4+ liquid from the feed gas to so maintain a relatively lean gas to the downstream unit. Preferred plants further include a turboexpander that receives at least part of the C4+ depleted vapor phase, and a fractionator receiving a C4+ liquid from a phase separator. It is still further preferred that the fractionator receives the C2− vapor (methane- and ethane-rich vapor) from the lower section of the fractionator as a reflux. Notably, reflux used in prior arts is typically a liquid phase while the methane and ethane vapor in the contemplated matters is a two phase stream, Most preferably, the lower section of the fractionator operates at a higher pressure (e.g., at least 5 psi, more typically at least 20 psi) than the upper section.

One exemplary plant configuration is depicted in FIG. 5, in which wet feed gas 1 at a pressure of about 1170 psig and a temperature of about 96° F., having a typical composition as shown in the table of FIG. 6, is dried in a molecular sieve drier 51, forming stream 2. The so dried gas stream 2 is cooled to a temperature of about 12° F. in exchanger 52, forming stream 3, utilizing the refrigeration content from residue gas stream 16 and liquid stream 6. The so chilled gas stream 3 is then separated in phase separator 53 into a liquid portion, stream 5, and a vapor portion, stream 4.

The liquid portion 5 is letdown in pressure via JT valve 54 to a pressure of about 510 psig, chilled to about −16° F. forming stream 6, which is heated in exchanger 52 to about 82° F. prior to entering as stream 7 to the lower section 71 (e.g., within the first five feed trays) of fractionator 70. The vapor portion 4 is expanded via the turboexpander 55 to about 500 psig at about −57° F. to form stream 8, which is fed to the upper section 72 (e.g., within the first three feed trays) of the fractionator 70. As used herein, the term “about” in conjunction with a numeral refers to a range of that numeral starting from 20% below the absolute of the numeral to 20% above the absolute of the numeral, inclusive. For example, the term “about −150° F.” refers to a range of −120° F. to −180° F., and the term “about 1500 psig” refers to a range of 1200 psig to 1800 psig. Moreover, and unless the context dictates the contrary, all ranges set forth herein should be interpreted as being inclusive of their endpoints, and open-ended ranges should be interpreted to include commercially practical values. Similarly, all lists of values should be considered as inclusive of intermediate values unless the context indicates the contrary.

The operating pressure of the fractionator 70 is in the range of about 450 to about 550 psig, and the upper section temperature is in the range of about −65° F. to about −55° F., and the lower section in the range of about 25° F. to about 300° F. The overhead gas stream 14 is the residue gas with a methane content of about 85 mol %. It should be appreciated that the fractionator 70 operates at two different pressures for the two sections, with the lower section 71 operating at at least 5 psig, more typically 10 psig, even more typically 20 psig, and most typically at least 30 psig higher than the upper section 72. To generate such pressure differential between the upper and lower sections, a booster pump 61 pumps liquid stream 10, typically from a tray below the expander inlet stream 8, to form stream 11 that is fed to the lower section 71 of the fractionator.

The lower section 71 acts as a stripper using reboiler 62 that maintains the ethane content in the bottom liquid stream 12 to no more than 2 mole %, that is required to meet the vapor pressure specification of the LPG product. The lower section 71 produces a vapor side stream 9 that contains about 5 to 6 mol % of the propane and heavier components. The side stream is cooled in exchanger 63 to a temperature of about −50° F. forming stream 9 using refrigeration content in the fractionator overhead stream 14 (from the upper section). The majority of the propane is so condensed, and the two phase mixture is letdown in pressure to about 480 psig via JT valve 60 and fed to the upper section 72 of the fractionator 70 as a top reflux, which is in a two phase regime. This reflux configuration is different than heretofore known configurations and methods which require the reflux to be in a single liquid phase. The two phase reflux configuration avoids the complexity of additional reflux drum and reflux pump, and is a more efficient approach since external refrigeration is not required.

The residual refrigeration content in residue gas 16 is recovered in exchanger 52 by cooling the feed gas. The warmed residue gas stream 17 at about 88° F. is compressed to a pressure of about 670 psig by compressor 56 driven by the turbo expander 55, forming stream 18. Stream 18 is further compressed to about 1200 psig by the residue gas compressor 57, forming stream 19 which is cooled by air cooler 58 prior to being delivered to the sales gas pipeline as stream 20.

With respect to the feed gas it is generally contemplated that suitable feed gases will include C1, C2 and C3+, and may further comprise N2 and CO2. Consequently, it should be appreciated that the nature of the feed gas may vary considerably, and all feed gases in plants are considered suitable feed gases so long as they comprise C1 and C3 components, and more typically C1 to C5 and heavier components, and most typically C1 to C6 and heavier components. Therefore, particularly preferred feed gases include natural gas, refinery gas, and synthetic gas streams obtained from other hydrocarbon materials such as coal, crude oil, naphtha, oil shale, tar sands, and lignite. Suitable gases may also contain relatively lesser amounts of heavier hydrocarbons such as propane, butanes, pentanes and the like, as well as hydrogen, nitrogen, carbon dioxide and other gases. Depending on the particular feed gas, the pressure of the feed gas may vary. However, it is generally preferred that the feed gas has a pressure between about 700 psig to about 1400 psig, and more typically between about 1000 psig to about 1400 psig.

With respect to most suitable applications, contemplated configurations and methods use a single fractionator to recover at least 98% of the C4 and heavier hydrocarbons, and 60% to 80% of the C3 component, without the use of external refrigeration. Therefore, it should be noted that feed gas cooling and/or cooling of the vapor product of the lower section are performed without use of external refrigeration. It should also be recognized that while a single column configuration is generally preferred, two separate columns with functions corresponding to the upper and lower sections are also deemed suitable for use herein. It is further contemplated that the dryer, separator, fractionator, heat exchanger, JT-valves, residue gas compressor, and turboexpander used in present configurations and methods are conventional devices well known to the skilled artisan.

Among other advantages of contemplated configurations, it should be particularly recognized that the separator produces a C5+ enriched liquid and a C5+ depleted vapor from a feed gas. Thus, so produced C5 enriched liquids may advantageously be fractionated in the lower section of the fractionator to meet the vapor pressure specification. Still further, it should be noted that the liquid is drawn from the upper section and is pressurized by a pump that allows the lower section to operate at a higher pressure than the upper section, and thus allows to provide cooling via a JT valve that produces a reflux to the column, while vapor regenerated from the lower section is cooled by the residue gas, providing a two phase reflux stream to the upper section, without the need to fully condense the reflux stream.

Additionally, it should be recognized that by using a feed cooler and feed separator, and further cooling of the vapors from the feed cooler and separation of the cooled vapors in the separator (to form a C5+ enriched liquid and a C5+ depleted vapor) most, if not all of the heavier components are removed from the feed gas. Consequently, the composition of the material flowing through the cold section is substantially stabilized as processing of heavy components in the feed gas in the upper section of the fractionator can be eliminated. Therefore, the heat duties, the turbo expander, and the fractionator will operate at the most efficient points. Thus, contemplated configurations and processes allow handling of a rich feed gas composition, thereby eliminating the complexity of a refrigeration unit of most prior arts. Viewed from another perspective, contemplated processes maintain constant operating conditions for the NGL recovery plant by removal of the C5+ components in the feed gas. According to previously performed calculations (data not shown), contemplated configurations will achieve at least 60%, and more typically 78% propane recovery, and at least 95%, and more typically 98% butane recovery (see FIG. 6). Further contemplations, configurations, and methods suitable for use herein are described in U.S. Pat. Nos. 6,601,406, 6,837,7070, 7,051,552, 7,051,552 and 7,377,127, all of which are incorporated by reference herein.

It should be apparent to those skilled in the art that many more modifications besides those already described are possible without departing from the inventive concepts herein. The inventive subject matter, therefore, is not to be restricted except in the spirit of the appended claims. Moreover, in interpreting both the specification and the claims, all terms should be interpreted in the broadest possible manner consistent with the context. In particular, the terms “comprises” and “comprising” should be interpreted as referring to elements, components, or steps in a non-exclusive manner, indicating that the referenced elements, components, or steps may be present, or utilized, or combined with other elements, components, or steps that are not expressly referenced. Where the specification claims refers to at least one of something selected from the group consisting of A, B, C . . . and N, the text should be interpreted as requiring only one element from the group, not A plus N, or B plus N, etc. 

What is claimed is:
 1. A method of hydrocarbon dew point adjustment of a natural gas comprising C3 and C4 and heavier components, comprising: cooling the feed gas in a feed gas exchanger using a liquid phase of the cooled feed gas and an overhead product of an upper section of a fractionator; separating the cooled feed gas in a phase separator into the liquid phase and a vapor phase; feeding the liquid phase into a lower section of a fractionator and feeding the vapor phase into the upper section of the fractionator; wherein the upper and lower sections of the fractionator are coupled to each other such that (a) an expansion device reduces pressure of and provides a vapor product of the lower section to the upper section as a two phase reflux stream, and (b) a pump increases pressure of and provides a liquid product of the upper section to the lower section; cooling the vapor product in an overhead exchanger using the overhead product of the upper section of the fractionator; and wherein the fractionator is operated such that recovery of the C3 from the feed gas is between 60% and 80%, and recovery of the C4 and heavier components from the feed gas is at least 95%.
 2. The method of claim 1 wherein the fractionator is operated at a pressure of between 450 to 550 psig, and wherein the upper section is operated at a pressure that is at least 10 psig lower than a pressure of the lower section.
 3. The method of claim 1 wherein the fractionator is operated at a pressure of between 450 to 550 psig, and wherein the upper section is operated at a pressure that is at least 30 psig lower than a pressure of the lower section.
 4. The method of claim 1 wherein the upper section is operated at a temperature of −65° F. to −55° F. and wherein the lower section is operated at a temperature of 25° F. to 300° F.
 5. The method of claim 1 further comprising a step of expanding the vapor phase in a turbo expander and reducing pressure of the liquid phase in a second expansion device before feeding the vapor phase and the liquid phase into the upper and lower sections of the fractionator, respectively.
 6. The method of claim 1 wherein at least one of feed gas cooling and cooling of the vapor product of the lower section is performed without use of external refrigeration.
 7. The method of claim 1 wherein feed gas cooling and cooling of the vapor product of the lower section to form the two phase reflux stream is performed without use of external refrigeration.
 8. A processing plant for hydrocarbon dew point of a natural gas feed gas delivered from a feed gas source, comprising: a feed gas exchanger fluidly coupled to the feed gas source and configured to cool the feed gas using a liquid phase of the cooled feed gas and an overhead product of an upper section of a fractionator; a phase separator fluidly coupled to the feed gas exchanger and configured to separate the cooled feed gas into the liquid phase and a vapor phase; wherein the fractionator further comprises a lower section that is configured to receive the liquid phase and the vapor phase in the upper section; wherein the upper and lower sections of the fractionator are coupled to each other such that (a) an expansion device reduces pressure of and provides a vapor product of the lower section to the upper section, and (b) a pump increases pressure of and provides a liquid product of the upper section to the lower section; and an overhead exchanger that is configured to cool the vapor product from the lower section to form a two phase reflux stream using the overhead product of the upper section of the fractionator.
 9. The processing plant of claim 8 further comprising a second expansion device that is configured to receive and reduce pressure of the liquid phase, and further comprising a turbo expander that is configured to receive and reduce pressure of the vapor phase.
 10. The processing plant of claim 8 wherein the pump increases the pressure of the liquid product in an amount of at least 10 psig.
 11. The processing plant of claim 8 wherein the fractionator is configured to operate at a pressure of between 450 to 550 psig.
 12. The processing plant of claim 8 wherein the upper section of the fractionator is configured to operate at a temperature of −65° F. to −55° F. and wherein the lower section of the fractionator is configured to operate at a temperature of 25° F. to 300° F.
 13. The processing plant of claim 8 wherein the feed gas exchanger, the fractionator, and the expansion device are configured to allow for between 60% and 80% C3 recovery without use of external refrigeration.
 14. The processing plant of claim 8 wherein the upper section and the lower section of the fractionator are configured to operate at a pressure differential of at least 10 psig.
 15. A method of hydrocarbon dew point adjustment of a natural gas comprising C3 and C4 and heavier components, comprising: separating the feed gas into a liquid phase and a vapor phase; feeding the vapor phase into an upper section of a fractionator, and feeding the liquid phase into an lower section of the fractionator, wherein the upper section is operated at a lower pressure than the lower section; cooling a vapor product of the lower section to form a two phase reflux stream and feeding the two phase reflux stream into the upper section; increasing pressure in and feeding a liquid product of the upper section to the lower section; and withdrawing an overhead product of the upper section of the fractionator as a dew point adjusted natural gas, and withdrawing a bottom product of the lower section comprising between 60% and 80% of C3 and at least 95% of the C4 and heavier components of the feed gas.
 16. The method of claim 15 wherein the natural gas has a pressure of between 1000 psig and 1400 psig, and wherein the fractionator is operated at a pressure of between 450 to 550 psig.
 17. The method of claim 15 wherein a pressure difference between the upper section and the lower section of the fractionator is at least 10 psig.
 18. The method of claim 15 wherein the vapor product of the lower section is cooled using refrigeration content of the overhead product of the upper section of the fractionator.
 19. The method of claim 18 wherein the feed gas is cooled using refrigeration content of the overhead product of the upper section of the fractionator.
 20. The method of claim 18 wherein the feed gas is cooled using refrigeration content of the liquid phase after pressure reduction of the liquid phase. 